Solvent extraction and recovery of solvent



New. 15, 1955 M. R. FENSKE ET AL 2,

SOLVENT EXTRACTION AND RECOVERY OF SOLVENT Filed Nov. 12, 1952 2 Sheets-Sheet l WEE-E5 eier in mm \00 2. 2. mm. |l' I'll-I 6 B Tmmawt Qm B61 attornew obert 6.

)r'enerren lgFenske jnventors Nov. 15, 1955 M. R. FENSKE ET AL 2,723,940

SOLVENT EXTRACTION AND RECOVERY OF SOLVENT Filed Nov. 12, 1952 2 Sheets-Sheet 2 Qobcr'b 6. Ge iez" @z Pair/ Sw n, QCICOf'QGLCS United States Patent SOLVENT EXTRACTION AND RECOVERY OF SOLVENT Merrell R. Fenske and Robert G. Geier, State College,

Pa., assignors to Esso Research and Engineering Company, a corporation of Delaware Application November 12, 1952, Serial No. 320,016

14 Claims. (Cl. 196-13) This invention relates to solvent extraction and recovery and regeneration of the solvent by compression distillation. More specifically it relates to the recovery of solvent in processes wherein aromatic or polar type compounds are extracted from a hydrocarbon stream. In one of its more specific aspects the invention relates to the extraction of aromatics from catalytically cracked cycle stocks with mixed ammonia solvents wherein water injection is resorted to for the purpose of controlling solubility and wherein the solvent is recovered by compression distillation.

It is well known that catalytically cracked cycle stocks usually contain large amounts of saturates as well as aromatics, the latter chiefly of the polynuclear type. These latter aromatics are deleterious in subsequent catalytic cracking. However, the saturates and non-aromatic portions represent a valuable source of additional cracking stock and would be upgraded in value over their present use in the ordinary cycle stock if they could be economically isolated.

Liquid extraction is an eifective way to make this separation of saturates from aromatics or of the less polar types from the more polar or nucleated types. However, since a substantially pure solvent of properly controlled solvent power is normally essential for proper separation, solvent recovery costs and the cooling costs normally incident to solubility control in the extraction stage have heretofore greatly reduced the economic advantage of such separation. For instance, where a pure solvent may require about theoretical extraction stages to produce a rafiinate containing weight percent aromatics and an extract containing 90 weight percent aromatics from a cycle oil containing 50 weight percent aromatics, the same separation would require 7 such stages if the solvent itselfcontained 2 weight percent aromatics under otherwise identical conditions, and with 3 percent aromatics in the solvent even an infinite number of stages would not make a rafiinate containing only 10 percent aromatics.

It is the main object of the present invention to devise an improved extraction process. Another object is to reduce the solvent recovery costs of such a process without sacrificing the effectiveness of the separation. These objects, as well as the nature of the invention, will become more clearly apparent from the subsequent description and accompanying drawing.

In the drawing- Figure 1 represents a schematic flow plan illustrating a suitable arrangement of equipment particularly adapted for the objects of this invention, using mechanical compression as a means for supplying the heat required in the distillation stage.

2,723,940 Patented Nov. 15, 1955 ICC Figure 2 represents a similar flow plan wherein thermocompression is used instead of mechanical compression.

Unless indicated otherwise, it is to be understood that all percentages and ratios given herein are expressed on a weight basis.

In accordance with the present invention, an unusually effective and economical process for extraction of aromatics and highly polar compounds has been developed. In particular, it has been found that solvents which contain ammonia in mixture with certain other low molecular weight compounds are most satisfactory for extracting aromatics from mixed hydrocarbon streams. At the same time these solvents ofier an unusual advantage in that, unlike heavier solvents such as phenol, furfural, or aniline, they can be recovered from the raw aromatic extract by compression distillation.

In other words, the solvents suitable herein are such that the solvent vapor produced in the evaporation of the raw aromatic extract can be appreciably compressed in an essentially adiabatic process without condensation. Thus a vapor stream is obtained which is materially hotter than the aromatic extract being distilled. This hot vapor can, therefore, be returnd to the distillation column and there passed in indirect heat exchange with the raw relatively cool extract solution. This causes the vapor to condense and the liberated latent heat is used to evaporate additional quantities of solvent from the raw extract. Thus, unlike in conventional distillation, the heat required for evaporation of the solvent is constantly being reclaimed and utilized within the process. At the same time no cooling water is needed for condensing the distillate since the extract being fed into the distillation column serves this purpose.

This independence on cooling water brings about another important advantage. Whereas conventional distillation normally must be operated at 120 F. or higher, that is, at a temperature at least 20 to 40 F. higher than the temperature of the available cooling water, a comparable compression distillation step is not limited in any such way. Hence compression distillation can be efficiently operated at F. or lower.

The suitability of a solvent is determined by whether adiabatic compression of its saturated vapor produces a superheated (dry) vapor or a wet vapor. If a superheated vapor results, the solvent is operable. The production of a wet vapor on compression makes the solvent unsatisfactory. A test of applicability can be made by substituting in the following equation and solving for T3.

log Tl T1=saturation temperature at P1, K.

T2=saturation temperature at P2, K.

Ts=temperature produced by adiabatic compression from T1 and T2 to P2, K.

AH=heat of vaporization, calories per mol.

K=ratio of molal specific heat at constant pressure to molal specific heat at constant volume.

If T3 is higher than T2, the solvent is applicable for use in our process Such solvents include, among others, ammonia, methylamine, dimethylamine, ethylamine, propane, propylene, water, sulfur dioxide, carbon dioxide, mcthylchloride, and dichloro difluoromethane, and mixtures thereof. However, if T3 is equal to T2, an unsuitable wet vapor may be produced by compression. Normal butane, isobutane, phenol and furfural are examples of unsuitable materials. Usually solvents containing more than about 12 atoms per molecule are not suitable for use in our process. In addition to the foregoing characteristics, a suitable solvent also must have a normal boiling point within a practical range, for instance, between about 50 and +100 F.

Another feature of this invention is related directly to the extraction stage of the process. Normally an extrac tion stage comprises an extraction or stripping section and an enriching section. When using ammonia solvents in accordance with the present invention, the extraction section is generally located below the feed inlet and the enriching section thereabove. At constant temperature such as usually prevails in the extraction or stripping section, the solubility or weight percent of oil dissolved in the solvent phase increases as the solvent moves up the extractor toward the enriching section, This is due to the increasing aromaticity of the oil or hydrocarbon phase. However, in the enriching section it is desirable to maintain the solubility reasonably constant, since it has a great eifect on the selectivity or relative solubility of the solvent. This is illustrated by the data of Table I which were obtained in extracting a typical cycle oil of about 40 percent aromatics content with a solvent composed of 45 percent methylamine and 55 percent ammonia.

Table I.--Efiect of oil concentration on selectivity Concentration of Oil in Solvent Phase, Percent f a g Exactly analogous to relative volatility alpha in distill-Litton.

To counteract the reduction of selectivity due to increasing solubility, which in turn is due to the increasing aromaticity of the extracted oil, the temperature is progressively decreased in conventional enriching sections so as to keep solubility constant. For instance, to keep solubility constant at 25 percent, a temperature gradient of about 40 F. maybe required between the oil feed percent of water to an ammonia-amine mixture has been found to reduce the solubility of hydrocarbons therein to the same extent as lowering the temperature to F. Accordingly, depending on the nature .of the hydrocarbon feed, solvent, and anti-solvent, up to 20 percent of anti solvent can be injected as needed for solubility control. That is, the water or other anti-solvent is added to the solvent at two or more successive points in the upper part or enriching section of the extractor. While water is the preferred anti-solvent in the specific process described herein, any of the other solvents previously described may be used as an anti-solvent or solvent modifier for ammonia or for whichever other main solvent may have been chosen. Besides these, numerous other liquids, such as methanol, formamide, ethylene 813K 01 and ethy ene diamine, may be used as anti-solvents, since the anti-solvent normally remains in the liquid rather than in the vapor state, and hence need not possess the particular compressibility characteristics previously specified.

The water injection technique as proposed herein is particularly advantageous since its efiect is to keep the solvent-to-oil ratio'lower at the raffinate end of the extractor than at the extract end. This is due to the fact that the water introduced into the enriching section is in a solvent solution obtained as a bottoms product from the solvent recovery step. The use of a lower solvent-to-oil ratio at the raflinate end is helpful in that less of the extract solvent has to be evaporated. The distillation of a smaller amount of solvent effects a saving in steam in addition to a saving in heat transfer area required for the solvent evaporation. 'In fact the use of an increasing solvent-to-oil ratio resulting from the water injection is perhaps the ideal way to handle a separation on an equilibrium diagram such as is encountered here, that is, one in which the selectivity of the solvent decreases as the aromatic concentration increases. The reason is that the higher reflux ratio is applied in the region of low selectivity, which is where it is most needed. Note, for instance, that selectivity of the ammonia-amine solvent decreases from a value of about 6 when the oil stream contains 10% aromatics to a value of about 2 when the oil stream contains 90% aromatics.

EXAMPLE 1 The invention will now be described with reference to a specific example dealing with the extraction of a cracked cycle stock with an ammonia-amine solvent in a system illustrated in Figure l of the drawing. The feed is a catalytically cracked cycle oil boiling from 550 to 800 F. at atmospheric pressure and containing about 44% aromatics. Its gravity is 176 API. The solvent is a mixture of 40% ammonia and monomethylamine. The ratio of solvent to feed oil is 2.7 at the raffinate end and 3.3 at the extract end, as described later. Water equivalent to 8% of the solvent is injected between the oil feed stage and the extract end to maintain solubility at 25% in the enriching section, that is, to maintain a saturated solution containing 25% of oil in solvent throughout the enriching section. The products are an extract and a raiiinate containing 90 and 10% aromatics, respec-' Referring to Figure l, the cycle oil is fed into extractor 10 through line 1 at a temperature of 120 F. and a pres sure of about 230 p. s. i. a. Ammonia-methylamine solvent at the same temperature and pressure is fed through line 31. Aliquid extract is removed through overhead line 2 and is composed of solvent and 25 oil. The

' oil portion of the extract contains aromatics. The

extract at F. and 230 p. s. i. a. is passed to compression distillation'column 20 where solvent is boiled off at a pressure of 210 p. s. i. a. and a vapor temperature of 120 'F., that is, substantially at extraction temperature and pressure. Enough solvent is taken from distillation unit 20 as vapor to produce a bottoms having an overall composition of 46 percent oil, 43 percent solvent, and 11 percent water. This solvent vapor contains about 45 percent ammonia and 55 percent methylamine. Substantially all of this vapor passes through-line 21 to the suction side of a mechanicalcompressor 22. Thiscomprcs sor advantageously is a rotary compressor driven by a steam turbine 28 which may be driven by steam at 1050 F. and 500 p. s' i. a. However, other types of compress rs may be used.

Table Il.- -Key to Fig. 1

[Solvent-to-oil ratio-=27 to 1 by wt. at raflinatc end; solvent compositlon== 60 wt. percent monomethylamine+40 wt. percent ammonia; 8 wt. percent water injected into solvent between oil feed stage and extract end; aromatic yield=87%; aromatic purity=90%; non-aromatic yield=92%; non-aromatic purlty=90%] Stream wt Wt. per- Wt. per- Flow, 1,000 lbs. Per Hour Wt. perg cent Arocent Am- Temp, Pressure, cent Oil Solvent maticsin monlaIn F. p. s.1 Numeral Composition Oil Solvent Solvent Oil Water Total 15 0 0 1. 9 1.9 210 132. 1 0 0 132. 1 400 132.1 0 0 132.1 400 132. 1 0 0 132. 1 Bottoms 210 67. 9 72. 6 17. 4 157. 9

Steam to Turbine 0 0 18. 4 18.4 Solvent Feed 179. 8 0 0 179. 8

Reflux Oil 75 25 90 25 145 210 14. 2 42. 6 0 56. 8 46- .do 75 25 90 25 120 210 14. 2 42. 6 0 56.8 52. Solvent Vapor- 0 100 30 120 170 32. 8 0 O 32. 8 54 Liquid Solvent 0 100 30 120 170 32.8 0 0 32. 8 57 Bottoms.- 39 31 90 20 190 170 18.5 22. 9 17. 4 58. 8

59 Steam from Turbine 280 O 0 10.13 10. 3 61- Water Layer 5 48 90 190 170 17.4 1. 7 l7. 4 36. 5 64- .do 120 170 17.4 1.7 17.4 36. 5 65- Oil La er 190 170 1.1 21.2 0 22.3 175 170 3. 5 28.3 0 31. 8

210 15 0 28. 3 0 28. 3 120 15 3. 5 0 6.0 9. 5 280 15 0 0 6.0 6.0 Solvent Vapor- 210 150 4. 9 0 0 4. 9 Liquid Solvent 210 150 4. 9 0 0 4.9

Flashed Raflinate. 93 7 10 30 210 150 2. 9 38. 3 0 41. 2 d0 93 7 10 30 340 150 2. 9 38. 3 0 41. 2 Steam 360 150 0 0 4. 3 4. 3 Rafiiinate+Water 360 150 0 38. 3 10. 3 48. 6 -.do 210 150 0 38.3 10. 3 48. 6

Liquid Solvent 0 100 30 120 150 9.6 0 0 9. 6 Reflux S0lvent 0 30 120 150 3. 2 0 O 3.2 Liquid Solvent 0 30 120 150 6. 4 0 0 6. 4 Ratfinate Product. 0 210 150 O 38.3 0 38. 3 Water 1 210 150 0 0 10. 3 10.3

In compressor 22 the solvent vapor is compressed to a pressure of about 400 p. s. i. a., whereby it is superheated to about 180 F. The compressed superheated vapor then passes through line 23 to reboiler 24 which is positioned in the bottom part of distillation tower 20. Thus the superheated vapor is cooled to a saturation temperature of about 165 F., at which point it is condensed with the liberation of heat. This heat is used for boiling ofi additional solvent in tower 20. The condensed vapor stream 25' is passed to storage tank for reuse. A small amount of condensate may also be returned through line 27 to serve as reflux in tower 20, to improvethe purity of the solvent distillate stream 21.

An auxiliary heat source such as steam coil 4 is provided to heat the initial raw extract to its boiling point. After that the compressor 22 is started and supplies further heat by mechanical means, whereas the auxiliary heat source is then cut out.

The bottoms from tower 20 contain 46% oil, 43% solvent, and 11% water. They have a boiling point of about 145 F. and form a two-phase emulsion at that temperature. A bottoms stream26 is, therefore, taken to settler for phase separation. The top or solvent layer will contain 25% oil, 56% solvent, and 19% water. The bottom layer will contain about 75% oil and 25% solvent. 1 r A The solvent layer is passed through line 41'to a conven tional distillation tower for further solvent removal.

Sufiicient solvent is removed in tower 50 to produce a bottoms stream 57 having an overall composition of about 39% oil, 31% solvent and 30% water. These bottoms exist as two liquid phases at their boiling point, 190 F. Thus after passage to settler 60, the bottoms separate into a water layer consisting of; 48% spl vent,

5% oil, and 47% water, and an oil layer containing 5% solvent and 95% oil. A solvent vapor stream 52 containing about 30% ammonia and 70% methylamine is removed from tower 50 and passed through water condenser 53. The resulting liquid solvent stream 54 is returned to storage tank 30. A small or appropriate amount of condensed solvent may be returned through line 55 as a reflux for tower 50. Heat for column 50 may be supplied by introducing steam through line 59. Since these bottoms boil at about 190 F., atmospheric exhaust steam from turbine 28 can be used since such steam has a temperature of 280 F.

While more solvent could be removed in compression distillation unit 20, this would materially increase the boiling point of the bottoms 26. Consequently, the vapor 21 would have to be compressed more to provide the required temperature differential across the reboiler surface, and this would usually reduce the outstanding advantages of compression distillation because of the increased work of compression and also because of the higher pressure equipment required. The amount of solvent removed in the compression distillation unit is usually fixed so that the temperature differential between the boiling bottoms and the condensing compressed vapor is at least 20 F. when an adiabatic compression ratio of about 2 to 1 is used.

The water phase from separator is recycled through line 61. After cooling to about F. in cooler 62, it is introduced through line 63 into the enriching section of extractor 10 at three or more injection points 64 located between the oil feed stage and the extract end. The amount of water thus injected 'into extractor 10 amounts to about 8% based on the solvent. The antisolvent efiect otthiswater serves to maintain the solusolvent is withdrawn from separator 60 through line 65 and combined with oil stream 43, which is part of the heavy phase 42 withdrawn from separator 40. Stream 43, which consists of about 75% oil and solvent is suflicient in amount to produce the desired amount of aromatic extract product when combined with stream 65. The remainder of the heavy phase 42 from separator is passed through line 44. After cooling to 120 F. in water cooler 45, it is returned through line 46 to the extract end of extractor 10 as reflux.

The combinedpro'duct in line 66 typically may consist of about 89% oil and 11% solvent, with only a trace of water. The combined product stream 66 is sent to tower 70 for'final solvent removal.

It will be noted that solvent feed stream 31, reflux stream 44, and water recycle stream 61, all have to be cooled to extraction temperature prior to their introduction into extractor 10. In the example this cooling is accomplished by heat exchange with cooling water in coolers 32, and 62, respectively. However, the required cooling can be obtained in other ways. For instance, a very advantageous simplification of equipment can be achieved by flashing all three streams at a low pressure which results in autorefrigeration. All of the resulting solvent vapor can then be condensed in condenser 73 which operates at atmospheric pressure.

The extract stream 66 entering tower is flashed to atmospheric pressure and stripped with open steam introduced through line 75. The amount of steam used for stripping is preferably at least that which will produce as top product a mixture composed of 40% solvent and The essentially non-aromatic raflinate leaves the bot toms of extractor 10 through line 11 at 230 p. s. i. a. and 120 F. It'contains 17% solvent and only a trace of water. This ratlinate stream is flashed at 150 p. s. i. a. and 210 F. in tower 80. The sensible and latent heat required for the flash may be supplid in heat exchanger 12 by steam 14 which may be at least in part exhaust steam from turbine 28. Additional extraneous steam may be supplied through line 15. The resulting solvent vapor stream 81 is condensed in condenser 82 and returned to storage tank30 through line 83.

The raflinate bottoms pass through lines 84 and 85 to finishing tower 90. This tower strips the remaining solvent from the raflinate'at 360 F. with open steam introduced at 91. ,Tower operating at 150 p. s. i. a. produces pure, water-free solvent as the distillate. The raffinate and water are taken as bottoms at about 360 F. through lines 92 and 94 and separated from each other in separator 100. The hot bottoms from tower 90 may be indirectly contacted in a heat exchanger 93 with the crude raftinate stream 84 to preheat the latter prior to flashing in tower 90. The water settled out in separator is discarded through line 102 while the rafiinate product is drawn ofi through line 101. This raflinate contains only 10% aromatics and represents a superior stock for catalytic cracking.

Tower 90 is also advantageously used to recover the solvent fromthe mixed solvent-water vapors produced in tower 70. For this purpose the distillate stream 72, after condensationin atmospheric condenser 73, is passed through line "74 to tower 90. The pure solvent vapor from tower-90 is condensed in condenser 95. A portion of the condensate stream 96 is returned through line 97 1 For turbine 28; exhaust used in towers 50 and 75, and in heater 12.

1 For ralhnate stripper 90. V

3 Additional steam for raflinate heater 12. r

The above total of about 25,000 pounds of steam per hour compares with about 80,000 pounds of steam per hour required for solvent recovery when conventional distillation is used instead of the described compression distillation in an otherwise similar system. Compression'distillation, of course, requires somewhat higher grade steam than conventional distillation. In addition to this saving on the weight of steam, compression distillation saves cooling Water. Finally, as also pointed out earlier herein, a major amount of cooling is also saved in the present invention by water injection, which makes it possible to 7 control solubility in the extractor without having to rely on any temperature differential. It is thus clear'that solvent extraction systems embodying compression distillation have major practical advantages in commercial operations, particularly when water injection is also used. The data show that water injection does not add to the steam requirements, nor does it make the solvent recovery more complex.

, While the solvent employed in the example consisted of 40% ammonia and 60% monomethyla mine, solvents containing the same ingredients in different proportions may often be preferred. Such changes in solvent composition can be used to adjust the solubility of the solvent to the desired value within a wide range of temperatures. This in turn allows considerable freedom in choosing the temperature at which theextraction is performed. For instance, a solvent containing 35% of the amine and 65% ammonia will produce 10% solubility at 160 F. when in equilibrium with a catalytic cycle oil containing 10% aromatics. In comparison, the same solubility is obtained at F. if the solvent contains 60% of the amine and 40% ammonia. This equivalence of the eflect of 25% amine concentration and 40 F. in temperature level has been found to be generally applicable to hydrocarbon extraction work.

The solvent composition, temperature, and water injection ratio are the'variables which can be changed to produce the desired solubilities in'the extractor.

in the liquid phase.

The top temperature of the compression distillation tower should be at least as high as the entering. feed stream. The pressure on the tower is determined so that While the above example related principally to the upgrading of a catalytic cycle stock into a high-quality feed stock for catalytic cracking by removal of aromatics therefrom, the invention may be similarly applied to other feed stocks and for other purposes. For instance, the

The pressure on the extractor is fixed to maintain the solvent present invention may be useful for separating naphthas into a largely non-aromatic raifinate and an extract rich in aromatics and consequently characterized by an improved octane number.

A thermally reformed, or cracked, naphtha boiling between about 130 and 400. F., and containing 25 vol. per cent olefins, 25 vol. per cent aromatics, and 50 vol. per cent saturates, with a research octane number of 71, was extracted at about 75 F. with a solvent composed of 80 wt. per cent ammonia and 20 wt. per cent monomethylamine. The extract portion, comprising 50% of the original naphtha, had a clear research octane number of 97, While the ralfinate rated 46 octane number. Over 80% of the extract was composed of aromatics and olefins.

Since lower molecular weight hydrocarbons are generally more soluble, the solvent contains more ammonia than that used in the previous example relating to catalytic cycle stock. The use of a higher ammonia concentration in the solvent would increase somewhat the pressures throughout the process.

The process as described by the drawing is also operable for the case of naphtha extractions. process, utilizing the same principles, can be advanta' geously used. Towers 50 and 70 and separators 40 and 60 in the extract solvent recovery system would be deleted. The bottoms from tower 20 would pass to a duplicate of tower 90. Open steam would strip the solvent out of the extract hydrocarbon. Pure solvent vapor would be the top product and a mixture of hydrocarbon and water would be the bottoms. The bottoms would pass to a phase separator wherein liquid water is separated from the liquid hydrocarbon. Enough ofthe hydrocarbon layer to satisfy the material balance would be taken as extract product. The remainder of the hydrocarbon layer would be sent back to the extractor as reflux. The water required for injection into the extractor would be taken from the water layer in the phase separator. The remaining water would be discarded. By increasing the compression ratio in the compression distillation step, thus allowing more solvent to be removed in that step, the economics would be as good as. those of the previous example.

EXAMPLE 3 Instead of distilling the solvent with the aid of mechanical compression as illustrated in Fig. 1, thermocompression may be used. Thermocompression expands high pressure vapor through a nozzle to obtain a high velocity gas. This high velocity gas then draws low pressure vapor into a mixing zone and the mixture is compressed by a There a portion of the solvent ,vapor is condensed with.

the aid of cooling coil 121 or equivalent cooling means, preferably in equilibrium with the uncondensed vapor so that the temperature of the liquefied solvent is not reduced appreciably below its condensation point. The vapor not condensed in condenser 120 is withdrawn through line 127 and serves as the low pressure ammonia for the thermocompression ejector 126. The liquefied solvent is passed from condenser 120 through line 122 to an ammonia boiler 123. High pressure ammonia vapor needed by the ejector is produced in boiler 123 with the aid of a steam coil 124 or the like. The mixed superheated vapor stream 23 leaving the ejector 126 has a pressure intermediate to that of the high and low pressure streams. The pressure of solvent vapor 23 preferably corresponds to a saturation temperature about 20 F. above the boiling point of the bottoms in tower 20, inasmuch as condensation of stream 23fiin'reboiler 24 An even simpler serves to supply the heat required for evaporation of additional ammonia solvent from the extract phase being fed into tower 20. Thus, in the scheme described it may be desired that stream 23 have a temperature of about 160 F. and a pressure of 286 p. s. i. a. The resulting condensed solvent 25 and the concentrated raw extract 26 may thereafter be treated substantially as described earlier in connection with Figure 1.

The actual temperatures and pressures of the various streams in the thermocompression step may of course vary considerably depending upon the temperature level of tower 20, on the ratio in which it is desired to combine the high and low pressure streams in the ejectortype jet pump, as well as on the overall efiiciency of the thermocompression step. This latter may advantageously be in the range of about 0.75 to 0.90 and depends on the efficiency of expansion, compression, and transfer of momentum. Characteristic values for these three types of efficiencies may be 0.98, 0.95, and 0. 85, respectively. The weight ratioof condensed to uncondensed solvent leaving condenser 120 may be between about 2:1. and 1:3. As a result, for instance, when the weight ratio of high and low pressure ammonia in the ejector is 1:1, it is desirable to raise the ammonia in steam boiler 123 to a temperature of about 290 F. and a pressure of about 710 p. s. i. a. In this arrangement about 3.1 pounds of ammonia are evaporated with steam per pound of steam required by the steam boiler, as compared with about 2 pounds of evaporation per pound of steam in conventional distillation. At the same time, while not quite so efficient, thermocompression has an advantage over mechanical compression since it requires lower grade steam and no mechanical compressor. Furthermore, when the ratio of high to low pressure ammonia in the ejector is reduced to 1:2, then 3.9 pounds of ammonia can be evaporated per pound of steam. However, the additional economy is accomplished by the use of high pressure ammonia at about 400 F. and 1500 p. s. i. a. as opposed to the 290 F. and 710 p. s. i. ammonia when a 1:1 ratio is used.

A further advantage can be obtained by combining thermocompression with compression distillation. In the above described thermocompression scheme some heat is wasted in the preparation of high pressure ammonia,

since ammonia gas is first condensed, only to be subsequently reheated in the steam boiler. Instead of this, it is possible to replace the condenser 120 and steam boiler 123, shown in Fig. 2, by a mechanical compressor, e. g. a steam-driven centrifugal compressor. In such a case, when the ejector 126 is operated at a 1:1 ratio of high to low pressure ammonia, stream 21 is simply split in half, one-half of the vapor passing directly to the ejector whereas the other half is mechanically compressed to 710 p. s. i. a. and a temperature of 290 F. prior to entering the ejector. When this is done by a centrifugal compressor driven by steam at 1050 F. and 500 p. s. i. a., 0.31 pound of steam is required to compress a pound of ammonia vapor to the desired pressure, assuming that the steam expands to atmospheric pressure and that the efficiency of the turbine and compressor is 90 and percent, respectively. Since two pounds of solvent are evaporated in tower 20 per pound of high pressure ammonia in line 125, about 6.5 pounds of ammonia solvent are evaporated per pound of steam. Again, the efficiency is not so high as if the entire compression is done mechanically as shown in Fig. 1. However, only one-half the amount of material has to be compressed in this combined mechanical-thermocompression case.

Still other modifications and variations of the invention may become apparent to those skilled in the art without departing from the scope and spirit of the appended claims.

We claim:

1. In a process for recovering a solvent from a liquid extract soluti0n'containing the solvent substantially sat" urated with a hydrocarbon oil having an atmospheric boiling point between about 175 and 1000 B, said solvent being capable of being compressed adiabatically from, a saturated vapor to a superheated vapor, the improvement which comprises passing the substantially saturated extract solution from an extraction zone substantially at extraction temperature and pressure toa distillation zone, evaporating solvent from said saturated solution in the distillation zone substantially at said extraction temperature and pressure so that the resulting residual liquid becomes a two-phase system containing a. solvent-rich phase and an oil-rich extract phase, withdrawing the resulting solvent vapor overhead from the distillation zone, compressing substantially all of the withdrawnsolvent vapor sufficiently to heat it at least F. above the temperature of said residual liquid present in the distillation zone, passing the compressed hot solvent vapor in indirect heat exchange with the liquid present in the distillation zone, whereby the compressedsole vent vapor is condensed and additional solvent is evaporated .from the extract solution, recovering the condensed solvent, withdrawing said residual two-phase liquid from the distillation zone, and mechanically separating said residual liquid into a solvent-rich phase and an oil-rich extract phase.

2. A process according to claim 1 wherein the solvent contains from 20 to 100% ammonia.

3 A process according to claim 1 wherein the solvent is a mixture 015.20 to 70% ammonia and 80 to mono-- methylamine.

4. A process according to claim 3 wherein the hydrocarbon is a catalytically cracked cycle oil boiling between about 400 and 1000 F. at atmospheric pressure and containing about 30 to 60% of aromatic hydrocarbons.

5. In a process for recovering a solvent composed of 40 to65% methylamine and 60 to ammonia from a liquid extract. solution comprising about 60 to 90 parts of solvent, substantially saturated with to 10 parts of a petroleum cycle oil fraction, the improvement which comprises passing the solution under pressure into a distillation'zone where the solution is maintained at its boiling point evaporating enough solvent so that the residual liquid becomes a two-phase system containing a solventrich phase and an oil-rich extract phase, withdrawing an overhead stream of saturated solvent vapor from the distillation zone, compressing substantially all of the withdrawn solvent vapor suthciently to heat it at least 20 F. above the aforesaid boiling point, passing the compressed hot solvent vapor in indirect heat exchange with the solution present in the distillation zone, recovering the resulting condensed solvent, withdrawing said two-phase liquid from the distillation zone, and separating said solvent-rich liquid phase from said oil-rich liquid extract phase.

. 6. A process according to claim S wherein the entire stream of saturated solvent vapor withdrawn from the distillation zone is compressed mechanically to a pressure corresponding to a saturation temperature at least about 20 F. above the boiling point of the said extract solution.

7. A process according to claim 5 wherein the saturated solvent vapor withdrawn from the distillation zone is separated into an ejector stream and a direct stream in a weight ratio ranging between about 2:1 and 1:3, the ejector stream is compressed to an elevated pressure and a superheated condition, and the two vapor streams are mixed to produce a superheated mixed vapor having a pressure corresponding to a saturation temperature at least about 20 F. above the boiling point of the said extract solution.

8. A process according to claim 7 wherein the ejector stream is compressed mechanically.

9. A process according to claim 5 which comprises the specific steps of withdrawing the saturated solvent vapor from the distillationzone at a temperature of about 80 ture at leastabout 20 F. above the boiling point of said extract solution.

10. A process according to claim 5. wherein saidseparated liquid extract phase hydrocarbon bottoms is passed to a stripping zone where additional solvent is stripped from the said extract phase at relatively low pressure.

11. A process for separating aromatic hydrocarbons from a cycle. oil which comprises passing the cycle oil into. an intermediate part of anextraction zone atan extraction temperature of about to 170 F., introducing a liquid ammonia solvent substantially at extraction temperature into a lower part of the extraction zone, maintaining the extraction zone under sufiicient pressure to keep the. solvent mixture in liquid phase, withdrawing a rafiinate stream containing an oil of lessthan 20% aromatics: content from the bottom part of the extraction zone, withdrawing under pressure an extract stream containing liquid solvent substantially saturated with oil of at'least aromatics content from the upper part of the extraction zone, passing the extract stream into'a 'distilla'tion: zone maintained at substantially said extraction temperature and at a pressure permitting evaporation the extract stream, removing overhead the resulting solvent vapor so as to form a two-phase liquid containing an oil-rich extract bottoms phase and a solvent-rich bottoms phase in the bottom of said distillation zone, superheating substantially all of the withdrawn solvent vapor by mechanical compression to a temperature at least 20 .F. higher than the aforesaid boiling point, passing the superheated solvent vapor in indirect heat exchange with the liquid extract present in the distillation zone, recovering the resulting condensed solvent for reuse, withdrawing said two-phase liquid from the distillation zone, separating said solvent-rich bottoms phase from said oil-rich extract b'ottoms' phase, stripping the separated oil-rich extract phase in an extract stripping zone with steam at substantially atmospheric pressure to strip out'additional solventand to produce a concentrated extract rich in aromatics and containing less than 0.5% solvent, recovering the flashed off solvent, and recovering said concentrated aromatic extract. 7

12. A process according to claim 11 wherein the railinate is steam stripped in a ratlinate stripping zone at a pressure .suflicient to maintain water in liquid phase, the overhead vapors from the extract stripping zone are pumped into the raflinate stripping zone, a solvent vapor overhead'is recovered from the raffinate stripping zone, and a bottoms cut containing an essentially solvent-free oil and Water is also recovered from the rafiinate stripping zone.

13. A process according to claim 11 wherein the solvent comprises a mixture of ammonia and methylamine, the extraction zone is maintained throughout at uniform temperature, a water stream is'injected into the upper 13 tion into an enriching section of the extraction zone, and the aromatic extract phase is recovered.

14. In a process according to claim 11 wherein solvent is recovered from a substantially saturated extract stream, the further improvement which comprises returning a portion of the separated oil-rich liquid extract bottoms phase as reflux to an enriching section of the ex traction zone from which the first mentioned liquid ex tract stream is being obtained for solvent recovery.

References Cited in the file of this patent UNITED STATES PATENTS Atkins Oct. 8, 1940 Bolt et al. Oct. 24, 1944 Sweeney et a1. Mar. 12, 1946 Cummings et al Mar. 12, 1946 Davis Dec. 27, 1949 Dickinson Nov. 4, 1952 

1. IN A PROCESS FOR RECOVERING A SOLVENT FROM A LIQUID EXTRACT SOLUTION CONTAINING THE SOLVENT SUBSTANTIALLY SATURATED WITH A HYDROCARBON OIL HAVING AN ATMOSPHERIC BOILING POINT BETWEEN ABOUT 175 AND 1000*F., SAID SOLVENT BEING CAPABLE OF BEING COMPRESSED ADIABATICALLY FROM A SATURATED VAPOR TO A SUPERHEATED VAPOR, THE IMPROVEMENT WHICH COMPRISES PASSING THE SUBSTANTIALLY SATURATED EXTRACT SOLUTION FROM AN EXTRACTION ZONE SUBSTANTIALLY AT EXTRACTION TEMPERATURE AND PRESSURE TO A DISTILLATION ZONE, EVAPORATING SOLVENT FROM SAID SATURATED SOLUTION IN THE DISTILLATION ZONE SUBSTANTIALLY AT SAID EXTRACTION TEMPERATURE AND PRESSURE SO THAT THE RESULTING RESIDUAL LIQUID BECOMES A TWO-PHASE SYSTEM CONTAINING A SOLVENT-RICH PHASE AND AN OIL-RICH EXTRACT PHASE, WITHDRAWING THE RESULTING SOLVENT VAPOR OVERHEAD FROM THE DISTILLATION ZONE, COMPRESSING SUBSTANTIALLY ALL OF THE WITHDRAWN SOLVENT VAPOR SUFFICIENTLY TO HEAT IT AT LEAST 20* F. F. ABOVE THE TEMPERATURE OF SAID RESIDUAL LIQUID PRESENT IN THE DISTILLATION ZONE, PASSING THE COMPRESSED HOT SOLVENT VAPOR IN INDIRECT HEAT EXCHANGE WITH THE LIQUID PRESENT IN THE DISTILLATION ZONE, WHEREBY THE COMPRESSED SOLVENT VAPOR IS CONDENSED AND ADDITIONAL SOLVENT IS EVAPORATED FROM THE EXTRACT SOLUTION, RECOVERING THE CONDENSED SOLVENT, WITHDRAWING SAID RESIDUAL TWO-PHASE LIQUID FROM THE DISTILLATION ZONE, AND MECHANICALLY SEPARATING SAID RESIDUAL LIQUID INTO A SOLVENT-RICH PHASE AND AN OIL-RICH EXTRACT PHASE. 